Process for preparing a paraffin product

ABSTRACT

The invention is a process for preparing a paraffin product from a carbonaceous feedstock comprising (a) partial oxidation of the carbonaceous feedstock to obtain a mixture comprising hydrogen and carbon monoxide, (b) performing a Fischer-Tropsch reaction using the mixture as obtained in step (a) and recovering an off-gas from the Fischer-Tropsch reaction and a paraffin product, (c) hydrogenating at least a part of the off-gas from the Fischer-Tropsch reaction using a steam/off-gas mol ratio in the range of between 0.7 and 1.5 and a catalyst comprising copper and zinc and/or a catalyst comprising copper and manganese, and (d) preparing a hydrogen comprising gas from at least a part of the off-gas from the Fischer-Tropsch reaction.

This application claims the benefit of European Patent Application No. 11196229.6, filed Dec. 30, 2011, the entire disclosure of which is hereby incorporated by reference.

The invention is directed to a process for preparing a paraffin product from a carbonaceous feedstock comprising the following steps, (a) partial oxidation of the carbonaceous feedstock to obtain a mixture comprising hydrogen and carbon monoxide, (b) performing a Fischer-Tropsch reaction using the mixture as obtained in step (a) and recovering an off-gas from the Fischer-Tropsch reaction.

Such a process is described in WO-A-03/035590 and in WO-A-03/035591. In these publications it is described to recirculate the off-gas, referred to therein as tail gas, back into the Fischer-Tropsch reactor of step (b) or into gasifiers of step (a). Gasifiers produce CO and H₂ for onward reaction in the Fischer-Tropsch reaction. The off-gas will comprise according to one of these publications water, CO₂, CH₄, N₂, unreacted syngas (H₂ and CO) and vapour hydrocarbon products.

Whilst some of the off-gas can be recirculated to the Fischer-Tropsch reactor, recirculation of all of the off-gas to the Fischer-Tropsch reactor causes CO₂, CH₄ and inerts to build up thus reducing the amount of hydrocarbons produced by the Fischer-Tropsch reactor. This is the case when either coal, biomass or natural gas is used as carbonaceous feedstock.

Feeding the off-gas back to the gasifiers (which produce the H₂ and CO mixture) results in problems relating to the relative ratio of H₂/CO. For example, coal gasifiers typically produce a H₂/CO ratio which is lower than the preferred ratio to perform step (b). Recirculation of off-gas to the gasifiers will even further reduce this ratio. In case natural gas is gasified, feeding the off-gas back to the gasifiers (which produce the H₂ and CO mixture) will also results in problems relating to the relative ratio of H₂/CO.

The third option disclosed in WO03/035590 is to use the off-gas as a fuel for power generation. However, in practice this may not consume all of the available off-gas.

The present invention discloses a process, which makes use of the off-gas in a more efficient manner. A general description of this process can be found in EP1860063A1. The present invention relates to an optimised process. Especially the process hydrogenating the off-gas from the Fischer-Tropsch reactor has been optimised so that a very stable catalyst performance is achieved. This is advantageous because optimal process conditions can now be maintained for the entire catalyst life time, and the catalyst life time has been increased significantly.

The present invention relates to a process for preparing a paraffin product from a carbonaceous feedstock comprising the following steps,

(a) partial oxidation of the carbonaceous feedstock to obtain a mixture comprising hydrogen and carbon monoxide, (b) performing a Fischer-Tropsch reaction using the mixture as obtained in step (a) and recovering an off-gas from the Fischer-Tropsch reaction and a paraffin product, (c) hydrogenating at least a part of the off-gas from the Fischer-Tropsch reaction using a steam/off-gas mol ratio in the range of between 0.5 and 1.5, preferably between 0.7 and 1.5, and a catalyst comprising copper and zinc and/or a catalyst comprising copper and manganese, and (d) preparing a hydrogen comprising gas from at least a part of the off-gas from the Fischer-Tropsch reaction.

Applicants found that by performing step (d) the off-gas may find use as feedstock to prepare a hydrogen comprising gas product. This is advantageous because in the process to be improved by the present invention additional hydrogen is required in order to either optimise the hydrogen to CO ratio as explained above and/or to further upgrade the products as obtained in step (b) by one or more hydroprocessing steps.

The carbonaceous feedstock used in step (a) may be coal, biomass or natural gas.

Step (d) may be any process, which can prepare a hydrogen containing mixture. The process (d) may be performed in a hydrogen manufacturing unit. Preferably the process (d) is catalyzed and more preferably it is a reforming process, for example the well known steam reforming processes, especially steam methane reformer (SMR), adiabatic steam reforming (ASR), fired steam reforming, and auto thermal steam reforming (ATR).

In case the carbonaceous feedstock used in step (a) is coal, the hydrogen manufacturing unit used in step (d) is an ASR. In case the carbonaceous feedstock used in step (a) is natural gas, the hydrogen manufacturing unit used in step (d) preferably is an SMR.

Step (d) and step (a) are separate steps resulting in separate gaseous products. The gaseous products as separately obtained may be combined after performing the separate steps.

The olefins and/or paraffins as present in the off-gas are first hydrogenated prior to performing step (d). This is performed in step (c). Step (c) may be a single hydrogenation step or two or more hydrogenation steps. When step (c) comprises two or more hydrogenation steps different catalysts may be used in the different hydrogenation steps.

Hydrogenation step (c) is very useful. For example, when an adiabatic steam reformer (ASR) is used for step (d) the temperature of the feed to the ASR is usually rather high (inlet temperature of around 500° C.) in order to compensate for the low activity of the catalysts used therein. At such temperatures, the presence of CO in the inlet end of the hydrogen manufacturing unit causes coking according to the Boudouard reaction below (1).

2CO→C+CO₂  (1)

Similarly, when an SMR is used as hydrogen manufacturing unit for step (d), the presence of CO at the inlet may cause that coke is formed at the inlet of the SMR.

Olefins and paraffins are also known for causing coking of catalyst(s) in the hydrogen manufacturing unit that is used in step (d). Carbon deposition or coking leads to hot spots on the catalyst and consequently reduces their activity. The hot spots are also formed on the reformer reactor tubes, and reduce their lifetime. The carbon deposits can be avoided or mitigated if olefins and preferably CO are removed from the off-gas stream.

Preferably a portion of at least the olefins within an off-gas stream is removed or converted before using the off-gas as a feed in step (d). In addition, the reaction preferably converts the carbon monoxide into methane and/or carbon dioxide, especially by reaction with water under the formation of carbon dioxide and hydrogen.

Preferably also other compounds are removed from the off-gas stream which can result in carbon deposition, for example CO, paraffins heavier than LPG and light naphtha. Thus preferably a portion of the olefins present in said off-gas are hydrogenated. More preferably the carbon monoxide present in said off-gas is removed or converted either prior to, simultaneous with or after the olefin hydrogenation step. Even more preferably substantially all of the carbon monoxide is converted or removed prior to being fed into the reactant side of a hydrogen manufacturing unit.

Typically the carbon monoxide is converted to a species which is not liable to cause carbon deposition, for example carbon dioxide or methane. The carbon monoxide is preferably not converted to a species which is liable to cause carbon deposition, such as carbon. Preferably a catalyst is used which combines olefin hydrogenation activity and CO shift activity. In that way olefins and carbon monoxide are removed, while additional hydrogen is made.

Suitable catalysts for step (c) are copper-containing catalysts. Preferably the catalyst comprises copper and zinc, or copper and manganese, or copper, zinc and manganese.

Before use the catalyst may comprise copper oxide. Such a catalyst may be activated by a reduction process to obtain metallic copper. Preferably the activated catalyst to be used in step (c) comprises metallic copper on a bulk of zinc, which may comprise metallic zinc and/or zinc oxide. Another preferred activated catalyst comprises metallic copper on a bulk of manganese, which may comprise metallic manganese and/or manganese oxide. Mixtures of these catalysts may be used.

Additionally or alternatively, catalysts comprising copper, zinc and manganese may be used.

Preferably the syngas produced by the hydrogen manufacturing unit has a sulphur content of below 1 ppm.

Hydrogenation step (c) may be performed in a single reactor comprising a copper catalyst. Preferably the hydrogenation step (c) is achieved in at least two reactors each comprising a copper catalyst. In case two reactors are used, one reactor can be reloaded with a fresh catalyst without stopping the operation of step (c).

Steam is added to the copper catalyst comprising reactor(s) for step (c). Hydrogenation is achieved using a steam/off-gas mol ratio in the range of between 0.7 and 1.5 and a catalyst comprising copper and zinc and/or a catalyst comprising copper and manganese in the one or more reactors for step (c). The temperature of the copper catalyst comprising reactor(s) for step (c) is preferably 200-300° C. at the inlet, more preferably 220-270° C. at the inlet.

In case two reactors comprising copper are used for step (c), the second reactor for step (c) may be provided in series with the first reactor for step (c) and the hydrogen manufacturing unit for step (d). The first and second reactor may be positioned in any order or combined but preferably the second reactor is downstream of the first reactor and preferably upstream of the hydrogen manufacturing unit. The reactors comprising copper may be used in lead/lag configuration.

In addition to the copper comprising catalyst, a pre-reforming catalyst may be used in step (c). The pre-reforming catalyst is used downstream of one or more copper comprising catalysts. For example, one reactor may be used in step (c), which is loaded with a pre-reforming catalyst in the bottom and thereupon a copper comprising catalyst. Or, two reactors may be used in step (c), each loaded with a pre-reforming catalyst in the bottom and thereupon a copper comprising catalyst. Alternatively, two reactors may be used in step (c), the first loaded with a copper comprising catalyst, and the second loaded with a pre-reforming catalyst in the bottom and thereupon a copper comprising catalyst. Or, two reactors may be used in step (c), the first loaded with a copper comprising catalyst and the second loaded with a pre-reforming catalyst.

Preferably steam is added to the reactor(s) comprising a pre-reforming catalyst. Pre-reforming catalysts preferably comprise nickel. A suitable catalyst comprises nickel on an alumina support. For example, a suitable pre-reforming catalyst comprises 40-60 wt % NiO on an alumina support, calculated on the total weight of the catalyst. Other options for pre-reforming catalysts include Pt, Ru, Rh, precious metals or combinations thereof. Preferably the inlet temperature to the pre-reforming catalyst is between 300 and 500° C., more preferably 330-400° C. For the pre-reformer catalyst steam is preferably used in a steam/dry gas molar ratio of 1.5 to 1.5.

The hydrogen manufacturing unit produces in step (d) a hydrogen comprising mixture. Usually the H₂/CO ratio of the hydrogen comprising mixture is 4:1 till 9:1. Preferably a portion or all of the hydrogen comprising mixture produced by the hydrogen manufacturing unit is used as part of the hydrogen carbon monoxide mixture feed in step (b). This may be effected by for example blending this mixture with the mixture as obtained in step (a) or by directly feeding this mixture to step (b). The purity of the hydrogen may be increased by know processes such as membrane separation, pressure swing absorbers (PSA) or combinations of a membrane unit followed by a PSA.

A portion of the optionally further purified hydrogen comprising mixture, particularly the hydrogen, as obtained in step (d) is preferably used to upgrade the paraffin product as obtained in step (b). More preferably said upgrading comprises hydrogenation, hydroisomerisation and/or hydrocracking, hydrodesulphurisation and catalytic dewaxing. Such upgrading processes as for example illustrated in WO-A-02/070629 in the context of a Fischer-Tropsch process.

If one requires even more hydrogen it is preferred to also use an additional hydrocarbon feedstock as feedstock in step (d). Such an additional hydrocarbon feedstock may be a methane comprising gas, LPG and naphtha. The LPG and naphtha may be derived from a mineral source or may be the LPG and/or naphtha products as isolated and obtained from the paraffin product as obtained in step (b) of the process of the present invention. Examples of methane comprising gasses may be refinery off-gas, coal bed methane or natural gas. Coal bed methane is preferred when the solid carbonaceous feedstock is coal because the coal bed methane is often found in the same location as the coal. The additional methane comprising gas may be subjected to the same hydrogenation type steps as described above if the gas comprises similar components, which require removal prior to feeding the gas to step (d).

In step (a) a carbonaceous feedstock is partially oxidated with an oxygen comprising gas. This is also referred to as gasification. The carbonaceous feedstock may be coal, biomass or natural gas.

The gasification in step (a) may be carried out by partially oxidating natural gas. The gasification in step (a) may be carried out by partially oxidating natural gas according to the shell gasification process (SGP) by partial oxidation of natural gas using pure oxygen. Partial oxidation of natural gas using pure oxygen may be operated at 1100 to 1700° C. Preferably partial oxidation of natural gas using pure oxygen is operated at 1300 to 1500° C. and pressures up to 70 bar. Another example of a process for partially oxidating natural gas is described in WO-A-9603345 where a mixture of carbon monoxide and hydrogen is prepared by partial oxidation of natural gas in a co-annular burner using 99.5% pure oxygen and optionally carbon dioxide as moderator gas and in the absence of a catalyst. A further example is described in WO2008006787. In the process of WO2008006787 partial oxidation on a methane comprising feed is performed using a multi-orifice burner provided with an arrangement of separate passages, wherein the gaseous hydrocarbon having at elevated temperature flows through a passage of the burner, an oxidiser gas flows through a separate passage of the burner and wherein the passage for gaseous hydrocarbon feed and the passage for oxidiser gas are separated by a passage through which a secondary gas flows, wherein the secondary gas comprises hydrogen, carbon monoxide and/or a hydrocarbon.

The gasification in step (a) may be carried out by partially combusting coal with a limited volume of oxygen at a temperature normally between 800° C. and 2000° C. in the absence of a catalyst. If a temperature of between 1050 and 2000° C. is employed, the product gas will contain very small amounts of gaseous side products such as condensable tars, phenols and hydrocarbons. Suitable coals include lignite, bituminous coal, sub-bituminous coal, anthracite coal, and brown coal. Lignites and bituminous coals are preferred. In order to achieve a more rapid and complete gasification, initial pulveriation of the coal is preferred. Particle size is preferably selected so that 70% of the solid coal feed can pass a 200 mesh sieve. The gasification is preferably carried out in the presence of oxygen and steam, the purity of the oxygen preferably being at least 90% by volume, nitrogen, carbon dioxide and argon being permissible as impurities. Substantially pure oxygen is preferred, such as prepared by an air separation unit (ASU). If the water content of the coal is too high, the coal is preferably dried before use. The atmosphere will be maintained reducing by the regulation of the weight ratio of the oxygen to moisture and ash free coal in the range of 0.6 to 11, preferably 0.8 to 1.0. The specific details of the procedures employed form no part of the invention, but those described in U.S. Pat. No. 4,350,103 and U.S. Pat. No. 4,458,607, incorporated herein by reference, may be employed. Although, in general, it is preferred that the ratio between oxygen and steam be selected so that from 0 to 0.3 parts by volume of steam is present in the reaction one per part by volume of oxygen, the invention is applicable to processes having substantially different ratios of oxygen to steam. The oxygen used is preferably heated before being contacted with the coal, preferably to a temperature of from about 200 to 500° C. Step (a) is preferably performed by partial oxidation of a dry mixture of coal particles and a carrier gas with oxygen in a membrane walled gasification reactor. Membrane wall reactors are known and for example described in US-A-2006/0076272. Preferably the hot mixture of hydrogen and carbon monoxide as obtained in the gasification reactor is cooled by direct contacting the hot gas with liquid water, also referred to as a water quench.

For coal-derived syngas the H₂/CO ratio of the gas mixture obtained in step (a) generally about or less than 1, and is commonly about 0.3-0.6. Such a ratio is suited for an iron catalyzed Fischer-Tropsch reaction. Because the low temperature cobalt catalysed Fischer-Tropsch reaction has a higher consumption ratio of between 2.0 and 2.1, additional hydrogen is needed. By conversion of part of the carbon monoxide as present in the gas mixture obtained in step (a) by means of the water gas shift reaction an increased amount of hydrogen is obtained thereby adjusting the H₂/CO ratio of the gas mixture to a value more suited for performing step (b). As explained above part of the hydrogen as prepared in step (d) may also be advantageously be used to modify the H₂/CO ratio of the gas mixture, thereby requiring less of the water gas shift reaction.

The catalytic water shift conversion reaction provides a hydrogen enriched, often highly enriched, syngas, possibly having a H₂/CO ratio above 3, more suitably above 5, preferably above 7, more preferably above 15, possibly 20 or even above. The water shift conversion reaction is well known in the art and is for example described in the earlier referred to WO-A-03035591. Generally, water, usually in the form of steam, is mixed with the syngas to form carbon dioxide and hydrogen. The catalyst used can be any of the known catalysts for such a reaction, including iron, chromium, copper and zinc. Copper on zinc oxide is a known shift catalyst. A very suitable source for the water required in the shift reaction is the product water produced in the Fischer-Tropsch reaction. Preferably this is the main source, e.g. at least 80% is derived from the Fischer-Tropsch process, preferably at least 90%, more preferably 100%. Thus the need of an external water source is minimised. Another preferred source of water is the quench water used to cool the hot gas in step (a) as described above.

When the gas mixture obtained in step (a) is coal-derived syngas, the desired ratio of hydrogen and carbon monoxide of the gas mixture to be used in step (b) is preferably controlled by passing only part of the gas obtained in step (a) over the catalytic water shift reaction as described above. In this manner one can target the desired ratio in an efficient manner, independent of the quality, that is the proportions of carbon and hydrogen, of the solid carbonaceous feedstock.

Especially when the gas mixture obtained in step (a) is coal-derived syngas, the mixture of hydrogen and carbon monoxide of step (a) may be passed through a carbon dioxide/hydrogen sulfide (CO₂/H₂S) removal system. This may also be performed when the gas mixture obtained in step (a) is natural gas-derived syngas. The removal system may involve one or more removal units. The CO₂/H₂S removal system preferably uses a physical solvent process, especially methanol or sulfolan, preferably methanol. This process is based on carbon dioxide and hydrogen sulfide being highly soluble under pressure in the solvent, and then being readily releasable from solution when the pressure is reduced as further discussed below. This high pressure system is preferred due to its efficiency, although other removal systems such as using amines are known.

It is preferred to remove at least 80 vol %, preferably at least 90 vol %, more preferably at least 95 vol % and at most 99.5 vol %, of the carbon dioxide present in the optionally catalytically shifted syngas stream. This avoids the build-up of inerts in the Fischer-Tropsch process.

On an industrial scale there are chiefly two categories of absorbent solvents, depending on the mechanism to absorb the acidic components: chemical solvents and physical solvents. Each solvent has its own advantages and disadvantages as to features as loading capacity, kinetics, regenerability, selectivity, stability, corrosivity, heat/cooling requirements etc.

Chemical solvents which have proved to be industrially useful are primary, secondary and/or tertiary amines derived alkanolamines. The most frequently used amines are derived from ethanolamine, especially monoethanol amine (MEA), diethanolamine (DEA), triethanolamine (TEA), diisopropanolamine (DIPA) and methyldiethanolamine (MDEA).

Physical solvents which have proved to be industrially suitable are cyclo-tetramethylenesulfone and its derivatives, aliphatic acid amides, N-methylpyrrolidone, N-alkylated pyrrolidones and the corresponding piperidones, methanol, ethanol and mixtures of dialkylethers of polyethylene glycols.

A well-known commercial process uses an aqueous mixture of a chemical solvent, especially DIPA and/or MDEA, and a physical solvent, especially cyclotetramethylene-sulfone. Such systems show good absorption capacity and good selectivity against moderate investment costs and operational costs. They perform very well at high pressures, especially between 20 and 90 bara.

The physical adsorption process useable in the present invention is well known to the man skilled in the art. Reference can be made to e.g. Perry, Chemical Engineerings' Handbook, Chapter 14, Gas Absorption. The absorption process useable in the present process is a physical process. Suitable solvents are well known to the man skilled in the art and are described in the literature. In the present process the liquid absorbent in the physical absorption process is suitably methanol, ethanol, acetone, dimethyl ether, methyl i-propyl ether, polyethylene glycol or xylene, preferably methanol. The physical absorption process is suitably carried out at low temperatures, preferably between −60° C. and 0° C., preferably between −30 and −10° C.

The physical absorption process is carried out by contacting the light products stream in a counter-current upward flow with the liquid absorbent. The absorption process is preferably carried out in a continuous mode, in which the liquid absorbent is regenerated. This regeneration process is well known to the man skilled in the art. The loaded liquid absorbent is suitably regenerated by pressure release (e.g. a flashing operation) and/or temperature increase (e.g. a distillation process). The regeneration is suitably carried out in two or more steps, preferably 3-10 steps, especially a combination of one or more flashing steps and a distillation step.

The regeneration of solvent from the process is also known in the art. Preferably, the present invention involves one integrated solvent regeneration tower.

The gas mixture of step (a) may also be passed over additional removal systems, guards or scrubbing units, either as back-up or support to the CO₂/H₂S removal system, or to assist in the reduction and/or removal of other contaminants such as HCN, NH₃, COS and H₂S, metals, carbonyls, hydrides or other trace contaminants.

The Fischer-Tropsch synthesis of step (b) is well known to those skilled in the art and involves synthesis of hydrocarbons from a gaseous mixture of hydrogen and carbon monoxide, by contacting that mixture at reaction conditions with a Fischer-Tropsch catalyst.

Products of the Fischer-Tropsch synthesis may range from methane to heavy paraffinic waxes. Preferably, the production of methane is minimised and a substantial portion of the hydrocarbons produced have a carbon chain length of a least 5 carbon atoms. Preferably, the amount of C5+ hydrocarbons is at least 60% by weight of the total product, more preferably, at least 70% by weight, even more preferably, at least 80% by weight, most preferably at least 85% by weight.

Fischer-Tropsch catalysts are known in the art, and typically include a Group VIII metal component, preferably cobalt, iron and/or ruthenium, more preferably cobalt. Typically, the catalysts comprise a catalyst carrier. The catalyst carrier is preferably porous, such as a porous inorganic refractory oxide, more preferably alumina, silica, titania, zirconia or mixtures thereof.

The optimum amount of catalytically active metal present on the carrier depends inter alia on the specific catalytically active metal. Typically, the amount of cobalt present in the catalyst may range from 1 to 100 parts by weight per 100 parts by weight of carrier material, preferably from 10 to 50 parts by weight per 100 parts by weight of carrier material.

The catalytically active metal may be present in the catalyst together with one or more metal promoters or co-catalysts. The promoters may be present as metals or as the metal oxide, depending upon the particular promoter concerned. Suitable promoters include oxides of metals from Groups IIA, IIIB, IVB, VB, VIIB and/or VIIB of the Periodic Table, oxides of the lanthanides and/or the actinides. Preferably, the catalyst comprises at least one of an element in Group IVB, VB and/or VIIB of the Periodic Table, in particular titanium, zirconium, manganese and/or vanadium. As an alternative or in addition to the metal oxide promoter, the catalyst may comprise a metal promoter selected from Groups VIIB and/or VIII of the Periodic Table. Preferred metal promoters include rhenium, platinum and palladium.

A most suitable catalyst comprises iron as this catalyst is suited for the lower hydrogen to carbon monoxide ratio as typically obtained in step (a). However by performing the process according to the present invention it also becomes possible to use cobalt based Fischer-Tropsch catalyst, which require a higher hydrogen to carbon monoxide ratio. A most suitable catalyst comprises cobalt as the catalytically active metal and zirconium as a promoter. Another most suitable catalyst comprises cobalt as the catalytically active metal and manganese and/or vanadium as a promoter.

The promoter, if present in the catalyst, is typically present in an amount of from 0.1 to 60 parts by weight per 100 parts by weight of carrier material. It will however be appreciated that the optimum amount of promoter may vary for the respective elements which act as promoter. If the catalyst comprises cobalt as the catalytically active metal and manganese and/or vanadium as promoter, the cobalt:(manganese+vanadium) atomic ratio is advantageously at least 12:1.

The Fischer-Tropsch synthesis is preferably carried out at a temperature in the range from 125 to 350° C., more preferably 175 to 275° C., most preferably 200 to 260° C. The pressure preferably ranges from 5 to 150 bar abs., more preferably from 5 to 80 bar abs.

Hydrogen and carbon monoxide (synthesis gas) is typically fed to the three-phase slurry reactor at a molar ratio in the range from 0.4 to 2.5. Preferably, the hydrogen to carbon monoxide molar ration is in the range from 1.0 to 2.5.

The gaseous hourly space velocity may very within wide ranges and is typically in the range from 1500 to 10000 Nl/l/h, preferably in the range from 2500 to 7500 Nl/l/h.

The Fischer-Tropsch synthesis is preferably carried out in multi-tubular reactor, or in a slurry phase regime or an ebullating bed regime, wherein the catalyst particles are kept in suspension by an upward superficial gas and/or liquid velocity.

It will be understood that the skilled person is capable to select the most appropriate conditions for a specific reactor configuration and reaction regime.

Preferably, the superficial gas velocity of the synthesis gas is in the range from 0.5 to 50 cm/sec, more preferably in the range from 5 to 35 cm/sec.

Typically, the superficial liquid velocity is kept in the range from 0.001 to 4.00 cm/sec, including liquid production. It will be appreciated that he preferred range may depend on the preferred mode of operation.

Experiments have been performed at optimal process conditions for hydrogenation step (c).

EXAMPLE 1

The off-gas comprised 12.2 vol % CO, 23.9 vol % CO₂, 6.5 vol % H₂, 25.4 vol % N₂, 30.5 vol % CH₄, and 1.1 vol % C₂H₄.

Steam was added to this dry gas. The steam/off-gas molar ratio was 1.2. Hydrogenation was performed using a catalyst comprising copper and zinc. The temperature at the inlet of the reactor was 220° C. The pressure was 400 psig.

The very stable catalyst performance is shown in Table 1.

TABLE 1 Outlet Time on stream (hrs) Composition Inlet 40 64 N2 (vol. %) 25.35 22.59 22.62 CO (vol. %) 12.25 0.29 0.30 CH4 (vol. %) 30.47 27.51 27.50 CO2 (vol. %) 23.94 31.98 32.07 H2 (vol. %) 6.48 15.83 15.82 C2H4 (vol. %) 1.14 0.01 0.01 C2H6 (vol. %) 0.00 1.03 1.03

EXAMPLE 2

The off-gas comprised 12.7 vol % CO, 25.4 vol % CO₂, 6.8 vol % H₂, 25.3 vol % N₂, 28.8 vol % CH₄, and 0.97 vol % C₂H₄.

Steam was added to this dry gas. The steam/off-gas molar ratio was 0.8. Hydrogenation was performed using a catalyst comprising copper and manganese. An isothermal reactor was used. The temperature at the inlet and at the outlet of the reactor was 250° C. The pressure was 400 psig.

The extremely stable catalyst performance is shown in Table 2.

TABLE 2 Outlet Time on stream (hrs) Composition Inlet 360 N2 (vol. %) 25.3 22.5 CO (vol. %) 12.7 0.115 CH4 (vol. %) 28.8 25.9 CO2 (vol. %) 25.4 34 H2 (vol. %) 6.85 16.6 C2H4 (vol. %) 0.968 0 C2H6 (vol. %) 0.00 0.965

In Example 2, the CO conversion was 99%, which was stable during the entire duration of the test (360 hours). The ethylene conversion was 100%, which was stable during the entire duration of the test (360 hours). 

What is claimed is:
 1. A process for preparing a paraffin product from a carbonaceous feedstock comprising the following steps, (a) partial oxidation of the carbonaceous feedstock to obtain a mixture comprising hydrogen and carbon monoxide, (b) performing a Fischer-Tropsch reaction using the mixture as obtained in step (a) and recovering an off-gas from the Fischer-Tropsch reaction and a paraffin product, (c) hydrogenating at least a part of the off-gas from the Fischer-Tropsch reaction using a steam/off-gas mol ratio in the range of between 0.7 and 1.5 and a catalyst comprising copper and zinc and/or a catalyst comprising copper and manganese, and (d) preparing a hydrogen comprising gas from at least a part of the off-gas from the Fischer-Tropsch reaction.
 2. The process according to claim 1, wherein the steam/off-gas mol ratio is in the range of between 0.8 and 1.2.
 3. The process as claimed in claim 1, wherein hydrogenation step (c) is performed at a temperature in the range of between 200 and 300° C., preferably between 220 and 270° C.
 4. The process as claimed in claim 1, wherein hydrogenation step (c) is performed using two or more reactors comprising a copper comprising catalyst.
 5. The process as claimed in claim 1, wherein additionally a pre-reformer catalyst is used in step (c), the pre-reformer catalyst preferably comprising nickel.
 6. The process as claimed in claim 5, wherein the inlet temperature used for the pre-reformer catalyst is in the range of between 350 and 500° C.
 7. The process as claimed in claim 5, wherein the steam/dry gas molar ratio for the pre-reformer catalyst is in the range of between 0.5 and 1.5. 